Black oil conversion for maximum gasoline production



1970 R. J. J. HAMBLIN f 5 I BLACK OIL CONVERSION FOR MAXIMUM GASOLINE PRODUCTION Filed Deg. 23. 1968 1s o q \0 O b N J6 N c M .9 Q 5 10 Q 2: I N q A N I V $2 9 S m a k N L N o A 85:- J 2 as v a w s m a M '2 g 9 2 t E L 55 g E E s L fig Q 1 8J1); -I: #4 2 A Q m 5 y N cc 1": IO 22 0 UV) 9 I Y 0 2; I. 5.2 0 E0 Q- 4-O- 0 o n om INVENTOR- Robert J. J. Hamb/in MKQAZLM ATTORNEYS United States Patent Ofice Patented Dec. 29, 1970 3,551,323 BLACK OIL CONVERSION FOR MAXIMUM GASOLINE PRODUCTION Robert J. J. Hamblin, Deerfield, Ill., assignor to Universal Oil Products Company, Des Plaines, 11]., a corporation of Delaware Filed Dec. 23, 1968, Ser. No. 786,025 Int. Cl. Gg 37/02 US Cl. 208-58 9 Claims ABSTRACT OF THE DISCLOSURE A multiple-stage catalytic process for maximizing the yield of desulfurized gasoline boiling range hydrocarbons from asphaltic charge stocks containing non-distillable hydrocarbonaceous material. The product efiiuent from the intial conversion zone is first separated into a principally liquid and a principally vaporous phase, the latter being withdrawn from the separation zone through a mass, or bed, of hydrogenation/desulfurization catalyst disposed therein.

APPLICABILITY OF INVENTION The invention described herein is adaptable to a process for the conversion of sulfurous, heavy hydrocarbonaceous material into lower-boiling hydrocarbon products of reduced sulfur concentration. More specifically, the present invention is directed toward processing asphaltene-containing, contaminated charge stocks, sometimes referred to in the art as black oils. In particular, the process encompassed by my invention affords the maximum production of desulfurized gasoline boiling range hydrocarbons from such material.

Petroleum crude oils, particularly the heavy oils extracted from tar sands, topped or reduced crudes, and vacuum residuum, etc., contain high molecular weight sulfurous compounds in exceedingly large quantities. In addition, these black oils contain excessive quantities of nitrogenous compounds, high molecular weight organometallic complexes, principally containing nickel and vanadium, and asphaltic material. An abundant supply of such hydrocarbonaceous material exists, most of which has a gravity less than 200 API at 60 F. This material is generally further characterized by a boiling range indicating that 10% by volume, and often up to 50.0%, boils above a temperature of about 1050 F.; these are considered to be non-distillables. The utilization of these high molecular weight black oils, as a sources of move valuable liquid hydrocarbon products, is virtually precluded by present-day refining techniques, due especially to the exceedingly high sulfur and asphaltic concentrations.

The process of the present invention is particularly directed toward the catalytic conversion of black oils into substantially sulfur-free distillable hydrocarbons boiling in the gasoline boiling range, and in volumetric yields of 100.0% or more. Specific examples of black oils, illustrative of those to which the present process is applicable, include a vacuum tower bottoms product having a gravity of 7.1 API at 60 F., containing 4.05% by Weight of sulfur and 23.7% by weight of asphaltics; a topped Middle-East Kuwait crude oil, having a gravity of 110 API at 60 F., containing 10.1% by weight of asphaltenes and 5.20% by weight of sulfur; and, a vacuum residuum having a gravity of 8.8 API at 60 F., containing 3.0% by weight of sulfur and 4300 p.p.m. of nitrogen and having a 20.0% volumetric distillation point of about 1055 F. The principal difliculty heretofore encountered resides in the lack of a suitable technique which would afford stability to many catalytic composites at the relatively severe conditions of operation, required to convert the non-distillables into lower-boiling products substantially free from sulfur. Catalyst instability also stems from the presence of the asphaltic material, which consists primarily of high molecular weight, non-distillable coke precursors, insoluble in light hydrocarbons such as pentane or heptane. Generally, the asphaltic material is dispersed within the crude oil, and when subjected to severe operating conditions, has the tendency to agglomerate and polymerize, whereby the conversion thereof to more valuable oil-soluble products becomes extremely difficult.

Heretofore in the field of catalytic processing of such hydrocarbonaceous material, two principal approaches have been advanced: liquid-phase hydrogenation and vapor-phase hydrocracking. In former types of process, liquid phase oil is passed upwardly, in admixture with hydrogen, into a fixed-fluidized catalyst bed, or slurry of subdivided catalyst. This type process is relatively ineflective with respect to asphaltics which are dispersed Within the charge, since the probability of effecting simultaneous contact between the catalyst particle, the hydorgen necessary for prevention of coke formation and the asphaltic molecule is remote. Furthermore, the retention of unconverted asphaltics, suspended in a free liquid phase oil for an extended period of time, results in additional flocculation and agglomeration. In addition, the efficiency of hydrogen to oil contact, obtainable by bubbling hydrogen through an extensive liquid body, is relatively low. Vapor phase hydrocracking, which requires a catalytic system, is precluded due to extreme catalyst deactivation as a result of the deposition of coke thereon. This type process also requires an attendant high capacity regeneration facilitiy in order to implement the process on a continuous basis.

OBJECTS AND EMBODIMENTS An object of the present invention is to provide an economically feasible catalytic process for converting black oils into distillable hydrocarbons of lower molecular weight and boiling range. A corollary objective is to maximize the production of substantially sulfur-free gasoline boiling hydrocarbons from black oils boiling above the normal gasoline boiling range.

Another object is to convert sulfur-contaminated hydrocarbon charge stocks, the greater proportion of which exhibits a boiling range above a temperature of 1050 E, into lower-boiling distillable hydrocarbon products significantly reduced in sulfur concentration.

Therefore, in a broad embodiment, the present invention involves a process for the conversion of a sulfurous, heavy hydrocanbonaceous charge stock, into lower-boiling, desulfurized hydrocarbon products, which process comprises the steps of: (a) heating said charge stock to a temperature of from about 500 F. to about 750 F., and reacting the heated charge with hydrogen in contact with a catalytic composite in a first reaction zone maintained under an imposed pressure above about 1000 p.s.i.g.; (b) separating the resulting reaction zone efiiuent in a first separation zone at substantially the same pressure imposed upon said first reaction zone, to provide a first vapor phase and a first liquid phase, withdrawing said first vapor phase through a mass of desulfurization catalyst disposed within said first separation zone; (0) separating the resulting desulfurized first vapor phase, in a second separation zone, at substantially the same pressure and at a lower temperature, to provide a second liquid phase and a hydrogen-rich second vapor phase; (d) separating at least a portion of said first liquid phase in a third separation zone at a substantially reduced pressure to concentrate and recover a residuum fraction, and to provide a third liquid phase; (e) reacting said third liquid phase with hydrogen, in a second reaction zone, at hydrocracking conditions, and in contact with a hydrocracking catalytic composite; and, (f) separating the resulting second reaction zone efiiuent in a fourth separating zone into a hydrogen-rich third vapor phase and a fourth liquid phase, and recycling at least a portion of said fourth liquid phase to combine with said third liquid phase, prior to reacting the same in said second reaction zone.

In a more specific embodiment, the present invention affords a process for the conversion of a hydrocarbonaceous charge stock of which at least about 10.0% by volume boils above a temperature of about 1050 F., and which contains more than about 1.0% by weight of sulfur, into lower-boiling, substantially sulfur-free hydrocarbon products, which process comprises the steps of: (a) heating said charge stock to a temperature in the range of from about 500 F. to about 750 F., and reacting the heated charge stock with hydrogen, in a first reaction zone in contact with a catalytic composite, and under an imposed pressure of from about 1000 p.s.i.g. to about 4000 p.s.i.g.; (b) separating the resulting first reaction zone efiluent, in a first separation zone, at substantially the same pressure imposed upon said first reaction zone and at a temperature of from 700 F. to about 800 F., to provide a first liquid phase and a first vapor phase, withdrawing said first vapor phase through a mass of desulfurization catalyst disposed in said first separation zone; separating the resulting desulfurized first vapor phase, in a second separation zone, at substantially the same pressure imposed upon said first separation zone and at a reduced temperature in the range of 60 F. to about 140 F., to provide a second liquid phase and a hydrogen-rich second vapor phase, recycling at least portion of said second vapor phase to combine with said charge stock; (d) recycling at least a portion of said first liquid phase to combine with said charge stock, and separating the remainder, in a third separation zone, at a substantially reduced pressure, to concentrate and recover an asphaltic residuum free from distillable hydrocarbons, and to provide a third liquid phase; (e) reacting said third liquid phase, in a second reaction zone, with hydrogen at a pressure greater than about 500 p.s.i.g. and a temperature above about 600 F., to convert sulfurous compounds to hydrogen sulfide and hydrocarbons, and to produce lower molecular weight, normally liquid hydrocarbons; (f) introducing at least a portion of the resulting second reaction zone effluent in admixture with a previously hydrocracked product, into a third reaction zone, at hydrocracking conditions selected to produce additional lower molecular weight hydrocarbons; (g) separating the resulting third reaction zone efiluent, in a fourth separation zone, at substantially the same pressure and a temperature of from about 60 F. to about 140 F., to provide a hydrogen-rich third vapor phase and a fourth principally liquid phase; and (h) recycling said third vapor phase to combine with said third liquid phase and combining said fourth liquid phase with said second liquid phase; and (i) recycling a portion of the mixture of said second and fourth liquid phases to combine with said second reaction zone efiluent as the aforesaid hydrocracked product and recovering the remainder as said lower-boiling, substantially sulfur-free hydrocarbon products.

Other embodiments of my invention reside in particular operating conditions and the use of specific internal recycle streams. The latter include recycle of the first hydrogen-rich gaseous phase to combine with the fresh charge stock prior to reacting the same in the conversion zone; this gaseous phase constitutes more than about 80.0 mol percent hydrogen. At least a part of the first liquid phase is diverted and combined with the charge and hyhydrogen mixture serving as a solvent stream to maintain the asphaltics dispersed and available to both hydrogen and catalyst in the first reaction zone. In another embodiment, a second portion of the first liquid phase is cooled and recycled to the inlet of the first separation zone to serve as a quench of the reaction zone efiluent such that the temperature in said first separation zone is maintained below a maximum level of about 800 F. The material entering the first separation zone is controlled at a temperature Within the preferred range of from about 700 F. to about 750 F., higher temperatures permitting heavier hydrocarbons to be carried over in the vapor phase.

Since the hot heavy oil from the conversion zone can give rise to serious emulsification problems as a result of the co-production of water, the hot separator is also employed to separate the heavy oil as a liquid phase from a vapor phase containing lighter hydrocarbons, hydrogen and Water. This hot separator is maintained at essentially the same pressure as the reaction zone and at essentially the temperature of the reaction zone effluent; as above set forth, in a preferred embodiment, the temperature is controlled in the range of about 700 F. to about 750 F.

In serving its previously described function, the hot separator provides a first principally vapor phase rich in hydrogen, and containing normally liquid hydrocarbons, the greater proportion of which boils below about 750 F. Since a considerable quantity of this normally liquid material is produced from the conversion of the sulfurcontaining, virgin asphaltics, the stream will continue to be contaminated by sulfurous compounds, albeit of lower molecular weight. This first vapor phase, as hereinafter set forth in describing the accompanying drawing, is further separated, at substantially the same pressure, but at a lower temperature in the range of 60 F. to F. in order to concentrate the hydrogen in a second vapor phase for recycle purposes. Normally liquid hydrocarbons are condensed in this cold separator and removed therefrom as a second principally liquid phase. In view of the fact that sulfurous compounds remain, this second liquid phase would usually be combined with other liquid streams free from non-distillables, the mixture serving as the feed to a hydrorefining, or desulfurization, zone designed to prepare a charge suitable for hydrocracking.

In accordance with my invention, the upper portion of the hot separator has disposed therein a finite mass, or bed, of desulfurization catalyst. While the precise quantity of catalyst will be dependent primarily upon the physical and chemical characteristics of the charge stock, and the particular separation being effected in the hot separator, the amount will be such that the liquid hourly space velocity of the hydrocarbons therethrough, as if actually in liquid phase, is in the range of from about 0.5 :1 to about 8.0:1. Distinctly advantageous is the fact that the normally liquid hydrocarbons are in vapor phase as they contact the catalyst. This affords substantially complete desulfurization with the result that the subsequently condensed stream can be subjected to product separation without additional catalytic treatment.

As hereinbefore stated, the catalytic composite disposed in the upper portion of the hot separator, through which the first vapor phase is withdrawn, is intended to have desulfurization activity. The degree of hydrocracking activity also possessed by the catalyst in the hot separator is usually dependent upon the individual refinery demands and the character of the charge stock. Where desulfurization takes precedence over the production of gasoline boiling range hydrocarbons, the hydrocracking activity is generally less than those instances where the naphtha yield and degree of desulfurization are of at least the same priority. For the former, I prefer a catalyst in which the alumina to silica weight ratio is as high as 9.0/1.0, and at least about 7.0/1.0; for the latter, an alumina/silica ratio as low as 1.5/1.0 is suitable. Suitable catalytic composites are those in which catalytically active components from Groups VI-B and VIII, and especially the iron-group, are combined with the alumina/ silica composite. The metallic components are selected from the group of chromium, molybdenum, tungsten,

iron, cobalt and nickel; in some situations, noble metals including platinum, palladium, iridium, osmium, ruthenium, and rhodium may be employed. Exemplary of especially preferred catalysts are: 2.0% by Weight of nickel and 16.0% by weight of molybdenum, combined with a composite of 88.0% alumina and 12.0% silica; 1.8% by weight of nickel and 16.5% by weight of molybdenum, combined with a composite of 90.0% alumina and 10.0% by weight of silica; 0.06% by weight of cobalt, 5.0% by weight of nickel and 10.0% by weight of molybdenum, combined with 63.0% alumina and 37.0% by weight of silica; and, 0.05% by weight of cobalt, 4.2% by weight of nickel and 11.3% by weight of molybdenum, combined with a composite of 88.0% alumina and 12.0% silica.

SUMMARY OF INVENTION Before summarizing my invention, several definitions are believed necessary in order that a clear understanding be afforded. In the present specification and the appended claims, the phrase pressure substantially the same as is intended to connote that pressure under which a succeeding vessel is maintained, is the same as an upstream vessel, allowing only for the pressure drop experienced as a result of the flow of fluids through the system. For example, where the conversion zone pressure, measured at the inlet thereof is 2650 p.s.i.g., the hot separator will function at about 2530 p.s.i.g. Similarly, the phrase temperature substantially the same as," is used to indicate that the only reduction, in temperature stems from normally experienced loss due to the flow of material from one piece of equipment to another or from conversion of sensible to latent heat by flashing.

When employed herein, the terms principally vaporous and principally liquid, are intended to allude to a particular separated stream, the greater proportion of which is either vapor-phase or liquid-phase. By way of further explanation, with respect to the third separation zone, two streams are separately withdrawn, both of which are referred to as liquid phases, notwithstanding that one is considerably lighter than the other. The lighter liquid phase, as hereinafter indicated, is primarily a vacuum gas oil fraction withdrawn as a liquid, raised to operating pressure and charged to the second reaction zone.

With respect to the untreated fresh charge stock, the phrase, non-distillables connotes hydrocarbonaceous material having a normal boiling point above 1050 F. This temperature might be considered self-limiting; as a practical matter, in the presence of native asphalts, premature thermal cracking would occur at high vacuum distillation temperatures corresponding to this normal boiling point. Following conversion and desulfurization, distillable hydrocarbons can be recovered having normal boiling points as high as 1150 F. since, at the same vacuum, the converted material has a higher thermal stability.

Conversion conditions are intended to be those conditions imposed upon the first conversion zone in order to convert a substantial portion of the black oil to distillable hydrocarbons. As will be readily noted by those skilled in the art of petroleum refining techniques the conversion conditions hereinafter enumerated are significantly less severe than those being currently commercially employed. The distinct economic advantages, over and above those inherent in producing the more valuable distillable hydrocarbons, will become immediately recognized. The conversion conditions, within this first conversion zone, are intended to include temperatures above about 500 F., with an upper limit of about 750 F., as measured at the inlet to the catalyst bed. Since the bulk of the reactions are exothermic, a temperature increase through the catalyst bed is experienced and the reaction zone effluent will be at a temperature higher than the inlet. In order that catalyst stability be preserved, it is preferred to control the inlet temperature such that the effluent temperature does not exceed about 900 F. Hydrogen is admixed with the black oil charge stock, by means of compressive recycle, in an amount of from about 5,000 to about 50,000 s.c.f./bbl., at the selected operating pressure, and preferably in an amount of from about 5,000 to about 15,000 s.c.f./bbl. The operating pressure will be greater than 1000 p.s.i.g., and generally in the range of about 1500 p.s.i.g. to about 4000 p.s.i.g. The black oil passes through the catalyst at a liquid hourly space velocity (defined as volumes of liquid hydrocarbon charge per hour, per volume of catalyst disposed in the reaction zone) of from about 0.25 to about 2.0.

As hereinbefore set forth, hydrogen is employed in admixture with the charge stock, and preferably in an amount of from about 5000 to about 15,000 s.c.f./bbl. The hydrogen-containing gas stream, sometimes designated a recycle hydrogen, since it is conveniently recycled externally of the conversion zone, fulfills a number of various functions: it serves as a hydrogenating agent, a heat carrier, and particularly a means for stripping converted material from the catalytic composite, thereby creating still more available catalytically active sites for the incoming, unconverted hydrocarbon charge stock. Since some hydrogenation will be effected, there will be a net consumption of hydrogen; to supplement this, hydrogen is added to the system from any suitable external source.

The catalytic composite disposed within the first conversion zone can be characterized as comprising a metallic component having hydrogenation activity, which component is composited with a refractory inorganic oxide carrier material. The precise composition and method of manufacturing the carrier material is not considered essential to the present process, although a silicious carrier, either amorphous or crystalline, such as 88.0% alumina and 12.0% silica, or 63.0% alumina and 37.0% silica, are generally preferred. Suitable metallic components having hydrogenation activity are those selected from the group consisting of the metals of Groups VI-B and VIII of the Periodic Table, as indicated in the Periodic Table of The Elements, E. H. Sargent & Co., 1964. Thus, the catalytic composite may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, palladium, iridium, osmium,

rhodium, ruthenium and mixtures thereof. The metallic' components of Group VI-B are preferably present in an amount within the range of about 1.0% to about 20.0% by weight, the iron-group metals in an amount within the range of about 0.2% to about 10.0% by weight, whereas the platinum-group metals are preferably present in an amount within the range of about 0.1% to about 5.0% by weight, all of which are calculated as if the components existed within the finished catalytic composite as the elemental metal.

The refractory inorganic oxide carrier material may comprise alumina, silica, zirconia, magnesia, titania, boria, strontia, hafnia, and mixtures of two or more including silica-alumina, alumina-silica-boron phosphate, silicazirconia, silica-magnesia, silica-titania, alumina-zirconia, alumina-magnesia, alumina-titania, magnesia-zirconia, titania-zirconia, magnesia-titania, silica-alumina-titania, silica-magnesia-zirconia, silica-alumina-boria, etc. It is preferred to utilize a carrier material containing at least a portion of silica, and preferably a composite of alumina, silica and 4.0% to 35.0% by weight of boron phosphate, with alumina being in the greater proportion.

Similarily, the catalytic composites disposed within the second and third reaction zones, will comprise metallic components from the metals of Groups VI-B and VIII of the Periodic Table, and compounds thereof, which components are composited with one or more of the refractory inorganic oxides previously set forth. However, since the catalytic composites in these two subsequent zones are intended to fulfill different functions, the second and third zones will, in most situations, make use of different catalyst formulations. The function of the first reaction zone is essentially two-fold: it serves to concentrate a residum fraction containing sulfur while simultaneously producing distillable hydrocarbons. The quantity of sulfur remaining in the distillable fraction will, of course, be dependent upon the characteristics of the fresh black oil charge stock. However, most of this remaining sulfur is concentrated in the middle-distillate and gas oil ranges, along with residual nitrogenous compounds. In order to produce maximum yields of gasoline, it is necessary to remove these residual contaminants.

The second reaction zone is employed for the primary purpose of effecting the virtually complete removal of the nitrogenous compounds by converting the same into ammonia and normally liquid hydrocarbons, while simultaneously converting most of the residual sulfurous compounds into hydrogen sulfide and liquid hydrocarbons. Generally speaking, the overall process is enhanced by virtue of the fact that some hydrogenation and hydrocracking of the higher-boiling middle-distillate and gas oil components takes place. While the product efllue-nt from this second reaction zone may be subjected to separation, to remove a lower-boiling gasoline fraction containing hydrocarbons boiling below a temperature of about 400 F. to about 450 F., the present process lends its readily to series flow with respect to the second and third reaction zones. The primary purpose of the third reaction zone is to maximize the conversion into lower-boiling, gasoline range products.

Since the functions to be served within the second and third reaction zones are different, most applications of the present invention will involve different operating conditions, as well as different catalysts. Thus, the charge to the second reaction zone, in admixture with hydrogen in an amount of about 1000 to about 10,000 s.c.f./bbl. of liquid hydrocarbon charge, is raised to a desired operating temperature within the range of about 600 F. to about 1000 F., prior to contacting the catalytic composite. The reactions are effected under an imposed pressure of about 100 pounds to about 4000 pounds per square inch, the hydrocarbon charge stock contacting the catalytic composite at a liquid hourly space velocity (LHSV) within the range of from about 0.5 to about 10.0. In addition to the effective clean-up of the hydrocarbon charge stock, a significant degree of hydrocarbon conversion occurs whereby the heavier molecular weight hydrocarbons, boiling at a temperature above about 700 F., and including the higher-boiling nitrogenous compounds, are converted, by highly selective cracking reactions into lower-boiling hydrocarbons, from which the nitrogen is more readily removed. The conversion reactions are such that very little normally gaseous hydrocarbons are produced.

The catalyst disposed within the second reaction zone serves a dual function; that is, the catalyst is non-sensitive to the presence of both nitrogenous compounds and sulfurous compounds, while at the same time is capable of effecting the destructive removal thereof, and, as hereinabove set forth, the conversion of at least a portion of those hydrocarbons boiling at a temperature above about 700 F. A catalyst comprising comparatively large quantities of molybdenum, calculated as the element, composited with the carrier material of silica and from about 60.0% to about 78.0% by weight of alumina, is very efiicient in carrying out the desired operation. A preferred catalytic composite, for utilization in this reaction zone, comprises from about 4.0% to about 45.0% by weight of molybdenum. In addition to minor amounts of nickel, from about 0.2% to about 10.0% by weight, like quantities of cobalt and/or iron may be employed in combination with the relatively large amounts of molybdenum.

The charge to the third reaction zone in admixture with hydrogen in an amount of from about 1000 to about 10,000 s.c.f./bbl., of liquid hydrocarbons, it at a temperature within the range of about 500 F. to about 750 F.

Due to the characteristics of the charge stock to the third reaction zone, the operating conditions within the same are relatively mild. The third reaction zone is maintained under an imposed pressure within the range of about 100 to about 4,000 p.s.i.g., and the rate of hydrocarbon charge will be within the range of from about 0.5 to about 15.0 liquid hourly space velocity. Catalytic composites which comprise at least one metallic component selected from Groups VI-B and VIII of the Periodic Table, and a composite of silica and from about 12.0% to about 30.0% by weight of alumina, constitute hydrocracking catalysts suitable for use in the conversion of the charge stock into lower-boiling hydrocarbon products. The total quantity of catalytically active metallic components is within the range of from 0.1% to about 20.0% by weight of the total catalyst. The Group VI-B metal, such as chromium, molybdenum, or tungsten, is usually present within the range of from about 0.5% to about 10.0% by weight of the catalyst. The Group VIII metals, which may be divided into two sub-groups, are present in an amount of from 0.1% to about 10.0% by weight of the total catalyst. When an iron sub-group metal such as iron, cobalt, or nickel, is employed, it is present in an amount of from about 0.2% to about 10.0% by weight, while, if a noble metal such as platinum, palladium, iridium, etc., is employed, it is present within an amount within the range of from about 0.1% to about 5.0% by weight. With certain charge stocks, and particular, desired ultimate product yields and distribution, it is advantageous to employ a metal-promoted crystalline aluminosilicate hydrocracking catalyst in the third reaction zone.

The total effluent from the third reaction zone is introduced into a high-pressure, low-temperature separation zone from which a hydrogen-rich gas stream is withdrawn and recycled. The normally liquid hydrocarbons, containing some light parafiinic hydrocarbons and butanes, are combined with normally liquid hydrocarbons resulting from the previously described second separation zone, and fractionated to remove therefrom those hydrocarbons boiling within the desired gasoline boiling range.

Other conditions and preferred operating techniques will be given in conjunction with the following description of the present process. In further describing this process, reference will be made to the accompanying figure which illustrates one specific embodiment. In the drawing, the embodiment is presented by means of a simplified flow diagram in which such details as pumps, instrumentation and controls, heat-exchange and heatrecovery circuits, valving, start-up lines and similar hardware have been omitted as being non-essential to an understanding of the techniques involved. The use of such miscellaneous appurtenances, to modify the process, are well within the purview of one skilled in the art.

DESCRIPTION OF DRAWING For the purpose of demonstrating the illustrated embodiment, the drawing will be described in connection with the conversion of a reduced Middle-East crude oil blend in a commercially-scaled unit. It is to be understood that the charge stock, stream compositions, operating conditions, design of fractionators, separators and the like, are exemplary only, and may be varied widely without departure from the spirit of my invention, the scope of which is defined by the appended claims. With reference now to the drawing, the charge stock, having the properties set forth in Table I, is introduced into the process via line 1.

TABLE I.-REDUCED CRUDE PROPERTIES Gravity, API 60 F. Sulfur, wt. percent 5.0 Nitrogen, p.p.m. 2800 Total metals, p.p.m. 103 Conradson carbon, wt. percent 13.1 Heptane-insolubles, wt. percent 9.63

9 TABLE l-Continu ed Initial boiling point, F. 50% distillation, F. 988 End boiling point, F. 1 1050 1 55.0 by volume distillable at 1050 F.

The unit is designed to process 35,000 bbL/day of this material at a combined liquid feed ratio of about 2.0:1. It is intended that the fresh feed be converted into maximum quantities of 350 F. end point gasoline which are substantially free from sulfur. The charge in line 1 is admixed with 35,000 bbl./day of recycle streams from lines 2 and 3, the source of which is hereafter described. Combined feed ratios in the range of about 1.2 to about 4.0, are well suited for use in this process. Based upon the fresh charge stock, a recycle hydrogen-rich gaseous phase, in line 4, in an amount of 10,000 s.c.f./ bbl., is admixed with the total liquid in line 1.

The resulting. mixture, following suitable heat-exchange with various hot effiuent streams not illustrated, continues through line 1 into heater 5 wherein the temperature is increased to a level such that the reactor 7 inlet temperature is about 675 F. The heated mixture leaves heaters via line 6, and enters first reaction zone 7 at an LHSV of 0.5, based upon fresh liquid feed only. Reaction zone 7 is maintained under an imposed pressure of about 2,900 p.s.i.g. The reaction product effluent, at a temperature of about 775 F., is withdrawn via line 8, and is passed thereby into hot separation zone 9 at a pressure of about 2850 p.s.i.g., the drop in pressure being due solely to fluid flow through the system. After its use as a heat-exchange medium, the eflluent is at a temperature of about 750 F. as it enters separator 9. A principally vaporous phase is withdrawn from the hot separator 9 through line 11, and a liquid phase is withdrawn via line 14. The first principally vaporous phase withdrawn through line 11 passes through a bed of desulfurization catalyst 10 disposed in the upper portion of the hot separator 9. In the instant illustration, the catalyst comprises 1.8% by weight of nickel and about 16.0% by weight of molybdenum, combined with a composite of 37.0% silica and 88. 0% alumina, by weight on a dry basis.

A portion of the liquid phase being withdrawn via line 14, about 35,000 bbL/day, is diverted through line 3 to combine with the fresh feed in line 1. The remaining portion continues through line 14 into vacuum flash column 15. This liquid phase comprises primarily 700 F.-plus hydrocarbonaceous material. A component analysis of the reactor 7 effiuent, exclusive of hydrogen and recycled material to the reaction zone, is presented in the following Table II:

TABLE II.REACTOR 7 PRODUCT EFFLUENT Following withdrawal of the vaporous phase through desulfurization catalyst 10, the vapors continue through line 11 and are further separated in cold separator 12, after being cooled, and condensed, to a temperature of about 110 F. This serves to further concentrate the hydrogen in a second vaporous phase for the purpose of recycle via line 4, to combine with the charge stock in line 1, and to provide a normally liquid phase in line 13, principally comprising butanes and heavier hydrocarbons. The latter, by virtue of having passed through catalyst bed 10, being substantially completely free from sulfurous and nitrogenous compounds, can continue through line 13 directly to fractionation system 26. If catalyst bed 10 were not included, the normally liquid phase in line 13 would contain nitrogenous and sulfurous compounds. In this situation, the portion boiling below 350 F. would not be suitable for inclusion in the ultimate product from fractionation, and the fraction boiling above 350 F. would not be a suitable feed stock for hydrocracking in the absence of additional treatment. Make-up hydrogen from any suitable external source, may be introduced into line 4, preferably upstream from conventional compressive means not shown in the drawing.

The liquid phase in line 14 is in part recycled via line 3 to combine with the fresh charge stock in line 1, the remainder continuing through line 14 into vacuum flash column 15. Vacuum column 15 functions at a pressure of about 25-60 mm. of Hg, absolute, and the liquid phase is introduced thereto at a temperature of about 730 F. Although the separation is shown as being effected in a single vacuum vessel, it may be suitable effected by an initial low pressure flash zone, at a pressure in the range of -200 p.s.i.g. followed by vacuum flashing of the liquid phase from the initial flash. In any event, the function of the hot flash system is to concentrate the unconverted asphaltics in a residuum fractionation free from distillable hydrocarbons, line 16, and to provide a third liquid phase principally comprising 350 F.-plus normally liquid hydrocarbons in line 17.

The liquid phase in line 17 is employed as a heat exchange medium in order to lower its temperature to about 650 F., as measured at the inlet of the catalyst disposed in reactor 19. In the instant illustration, the charge to reactor 19 is about 27,660 bbL/day, and is ad mixed with 10,000 s.c.f./bbl. of hydrogen from line 18. The LHSV (liquid hourly space velocity) through reactor 19 is 1.4, while the pressure is approximately 2100 p.s.i.g. Reactor 19 product eflluent, at a temperature of about 750 F., as measured at the outlet of the catalyst bed, is withdrawn by way of line 20, and admixed thereby with a hydrocarbon recycle stream in line 21. The source of the latter is hereafter set forth. The hydrocarbonaceous mixture continues through line 21 into reaction zone 22. The temperature of the mixture, as it initially contacts the catalyst in reactor 22, is about 700 F. The product efiluent, at a temperature of 750 F., is withdrawn by way of line 23, cooled to a temperature of about F. and introduced into high pressure separator 24.

A third, principally vaporous, hydrogen-rich gaseous phase is withdrawn by way of line 18 for use as recycle to reactor 19. Cold separator 24 serves as the focal point for pressure control of the portion of the process system encompassing reactors 19 and 22. Thus, the pressure imposed upon reactor 19, by way of compressive means not illustrated, is such that cold separator 24 functions at about 2000 p.s.i.g. Reactor 22 will function at some intermediate pressure as a result of pressure drop due to the flow of material through the system. As previously set forth, the principal function served by reactor 22 is hydrocracking the heavier components of the charge into gasoline boiling range material. As a consequence, there is a tendency toward a significant temperature rise through the catalyst bed, resulting from the exothermicity of the reactions being effected. Since optimum operation results when the temperature rise within the hydrocracking zone is limited to about 50 F., quench streams, preferably hydrogen diverted from line 18, may be introduced into reactor 22 at intermediate loci of the catalyst bed. The use of quench streams for temperature control is a conventional technique, and is not essential to my invention. Normally liquid hydrocarbon products are withdrawn from cold separator 24 by way of line 25, admixed with the liquid phase from cold separator 12, and introduced therewith into fractionation system 26 by way of line 13.

The product efiluent in line 23, exclusive of recycle hydrogen, being a composite of the reactions effected in reactors 19 and 22, has the component analysis shown in the following Table III:

TABLE III.REACTOR 22 EFFLUENT Wt. Vol.

Component percent percent Ammonia 0. 16

Hydrogen sulfide. 1.22

Methane 0. 23

Ethanm 0. 46

Propane 2. 81

Butanes 10.81

lentanes 7. 82

Hexanes 10. 36

Heptane350 40. 96

650 F.plus 38.10

A preferred processing scheme involves providing a com- 2 bined liquid feed ratio, to reactor 22, in the range of about 1.5:1 to about 2.0:1. In the process being illustrated, the combined feed ratio is 1.8:1. When there is an excess of 350 F.-plus material, such excess continues through line 2 to combine with the fresh feed black oil in line 1. When this situation arises, a like quantity of the hot separator bottoms liquid is backed-out of the recycle in line 3, in order to preserve the desired combined feed ratio to reactor 7.

Butanes and lighter, normally gaseous, materials are withdrawn from fractionator 26 as an overhead product in line 27. A C /C side out is removed via line 28, and the desired C -350 F. gasoline is removed by way of line 29. The overall product distribution, based upon the charge stock rate of 35,000 bbL/day, including light gaseous material, is presented in the following Table IV:

TABLE IV.OVER.ALL PRODUCT YIELD AND DISTRIBUTION Wt. Vol. Component percent percent Bbl/day Ammonia Hydrogen sulfide Metl Residuum It should be noted, from the data presented in Table IV, that the volumetric yield of normally liquid hydrocarbons, including butanes and excluding the unconverted asphaltic residuum, is 40,315 bbl./day, or 115.2% by volume, based upon fresh charge stock. Of further interest is the fact that the 7,4000 bbl./ day yield of butanes consists of 70.0% iso-butanes, while the 4,940 bbL/day of pentanes comprises 93.0% iso-pentanes. The C /C product has a gravity of 80.9 API and a Clear Research Octane Rating of 85. The heptane-350 F. gasoline fraction, having a gravity of 54.l API, has a Clear Octane Rating of 67 (85 with 3.0 cc. TEL) and comprises 47.0 volume percent paraffins, 35.0% naphthenes and 18.0% aromatics. As such, and considering its sulfur concentration is essentially nil (less than 0.1 ppm. by weight), this fraction constitutes an excellent charge stock for a subsequent catalytic reforming unit. These, and many other advantages will be readily apparent to those possessing expertise in the art of petroleum refining processing.

Although the mass, or bed of desnlfurization catalyst is shown as being disposed in the upper section of hot separator 9, it is clear that the important aspect is the vaporous state of the material passing therethrough, and not the precise location of the catalyst. Thus, the catalyst may be disposed in a separate zone installed between hot separator 9 and cold separator 12, or in two or more zones which can be easily manifolded to permit swing-bed operation. In many instances, the latter scheme would be preferred from the standpoint of permitting catalyst regeneration without the necessity for a complete shut-down of the process.

The foregoing specification indicates the method by which the present invention is effected, and the benefits afforded through the utilization thereof.

I claim as my invention:

1. A process for the conversion of a sulfurous, heavy hydrocarbonaceous charge stock, into lower-boiling, desulfurized products, which process comprises the steps of:

(a) heating said charge stock to a temperature of from about 500 F. to about 750 F. and reacting the heated charge stock with hydrogen in contact with a catalytic composite in a first reaction zone maintained under an imposed pressure above about 1000 p.s.i.g.;

(b) separating the resulting reaction zone efiluent in a first separation zone at substantially the same pressure imposed upon said first reaction zone, to provide a first vapor phase and a first liquid phase, withdrawing said first vapor phase through a mass of desulfurization catalyst disposed within said first separation zone;

(0) separating the resulting desulfurized first vapor phase, in a second separation zone, at substantially the same pressure and at a lower temperature, to provide a second liquid phase and a hydrogen-rich second vapor phase;

(d) separating at least a portion of said first liquid phase in a third separation zone at a substantially reduced pressure to concentrate and recover a residuum fraction, and to provide a third liquid phase;

(e) reacting said third liquid phase with hydrogen,

in a second reaction zone, at hydrocracking conditions, and in contact with a hydrocracking catalytic composite; and,

(f) separating the resulting second reaction zone effluent in a fourth separating zone into a hydrogenrich third vapor phase and a fourth liquid phase, and recycling at least a portion of said fourth liquid phase to combine with said third liquid phase.

2. The process of claim 1 further characterized in that said hydrogen-rich second vapor phase is recycled to combine with said charge stock.

3. The process of claim 1 further characterized in that said third separation zone is maintained at a pressure of from subatmospheric to about 200 p.s.i.g.

4. The process of claim 1 further characterized in that said second and fourth separation zones are maintained at a temperature within the range of from about 60 F. to about F. and at a pressure substantially the same as that imposed upon said first and second reaction zones respectively.

5. The process of claim 1 further characterized in that at least a portion of said first liquid phase is recycled to combine with said charge stock.

6. The process of claim 1 further characterized in that at least a portion of said fourth liquid phase is recycled to combine with said charge stock.

7. The process of claim 1 further characterized in that the combined liquid feed ratio to said first reaction zone is in the range of about 1.2:1 to about 4.0: 1.

8. The process of claim 1 further characterized in that said hydrogen-rich third vapor phase is recycled to said second reaction zone.

9. A process for the conversion of a hydrocarbonaceous charge stock, of which at least about 10.0% by volume boils above a temperature of about 1050 F., and

which contains more than about 1.0% by weight of sulfur, into lower-boiling, substantially sulfur-free hydrocarbon products, which process comprises the steps of:

about 500 p.s.i.g., and a temperature above about 600 F., to convert sulfurous compounds to hydrogen sulfide and hydrocarbons, and to produce lower (a) heating said charge stock to a temperature in the range of from 500 F. to about 750 F., and remolecular weight, normally liquid hydrocarbons; (f) introducing at least a portion of the resulting secacting the heated charge stock with hydrogen, in a ond reaction zone efiluent, in admixture with a prefirst reaction zone in contact with a catalytic comviously hydrocracked product, into a third reaction posite, and under an imposed pressure of from zone, at hydrocracking conditions selected to produce about 1000 p.s.i.g. to about 4000 p.s.i.g.; additional lower molecular weight hydrocarbons; (b) separating the resulting first reaction zone effiuent, (g) separating the resulting third reaction zone effiuent, in a first separation zone, at substantially the same in a fourth separation zone, at substantially the pressure imposed upon said first reaction zone and same pressure and a temperature of from about at a temperature of from 700 F. to about 800 F., 60 F. to about 140 F., to provide a hydrogento provide a first liquid phase and a first vapor phase, rich third vapor phase and a fourth principally withdrawing said first vapor phase through a mass 15 liquid phase; and, of desulfiurization catalyst disposed in said first (h) recycling said third vapor phase to combine with separation zone; said third liquid phase and combining said fourth (c) separating the resulting desulfurized first vapor liquid phase with said second liquid phase; and,

phase, in a second separation zone, at substan- (i) recycling a portion of the mixture of said second tially the same pressure imposed upon said first sepaand fourth liquid phases to combine with said second ration zone and at a reduced temperature in the reaction zone efiluent as the aforesaid hydro-cracked range of 60 F. to about 140- F., to provide a product and recovering the remainder as said lowersecond liquid phase and a hydrogen-rich second vapor boiling, substantially sulfur-free hydrocarbon prodphase, recycling at least a portion of said second vanet por phase to combine with said charge stock; References Cited UNITED STATES PATENTS 3,365,388 1/1968 Scott DELBERT E. GANTZ, Primary Examiner A. RIMENS, Assistant Examiner U. S. Cl. X.R. 

